Ammonia synthesis

ABSTRACT

Ammonia synthesis gas is made from a raw gas comprising hydrogen, carbon dioxide and medium boiling point gases including nitrogen in excess of the proportion required an ammonia synthesis gas, by a pressure swing adsorption process characterized by feeding to the absorbent a raw gas in which hydrogen and total medium boiling point gases are present in a volume ratio in the range 1.25 to 2.5, and the medium boiling point gases comprise nitrogen to the extent of at least 90% v/v on the total of such gases. Preferred ways of making the raw gas, of ensuring purity of the product gas and of recovering useful heat are described.

This is a continuation of application Ser. No. 032,576filed Apr. 1,1987, which was abandoned upon the filing hereof, which is acontinuation of Ser. No. 703,531, filed Feb. 20, 1985, now U.S. Pat. No.4,695,442.

This invention relates to the production of hydrogen and in particularto the production of purified ammonia synthesis gas from a raw gas.

Conventionally most ammonia synthesis gas is made by one of these tworoutes:

A. Steam reforming route:

(a) incomplete catalytic reaction of reformable hydrocarbon with steam;

(b) reaction of the product of (a) with air to introduce nitrogen andadequately to complete reaction of hydrocarbons;

(c) catalytic shift reaction with steam of carbon monoxide in theproduct of (b);

(d) removal of carbon dioxide in a regenerable absorbent liquid; and

(e) removal of residual carbon oxides by methanation; and

B. Partial oxidation route:

(a) separation of air to give liquid nitrogen and high-concentrationoxygen;

(b) partial oxidation of a carbonaceous feedstock with the oxygen andpossibly also steam to give a carbon monoxide rich gas;

(c) catalytic shift reaction with steam of carbon monoxide in theproduct of (b);

(d) removal of carbon dioxide in a regenerable absorbent liquid; and

(e) contacting the resulting gas with liquid nitrogen to condense outresidual carbon oxides and to introduce nitrogen.

Hereinafter the product of stage (c) in either route, that is, the gasafter shift but before complete carbon dioxide removal will be referredto as raw gas.

Recently processes capable of removing various gases from mixtures withhydrogen by pressure-swing selective adsorption have been developed andput into industrial use for producing pure hydrogen. In one such processthe further stage of producing ammonia synthesis gas by adding nitrogento the pure hydrogen has been proposed (European Chemical News 20 Oct.15, 1978, 39, 47). In other proposals (BE-A-885126, GB-A-2103199) thenitrogen has been introduced as a purge gas in regenerating theselective adsorbent. In either of such processes the nitrogen is derivedfrom some extraneous source such as air separation and the raw gas hasbeen made either by route B or by a modified route A without thenitrogen-introducing stage (b). It would in principle be more convenientif the nitrogen could be introduced as air; however, in the proposal onthese lines that is described in GB-A-2126573 there is only a lowpercentage (72.4%) recovery of hydrogen. In DE-A-3206153 such a processis proposed but is said to be impracticable unless a nitrogen-passingadsorbent is used.

We have now discovered from a study of the adsorption properties of thegases involved that over a narrow range of raw gas composition thecontent of medium boiling point gases especially nitrogen is in balancewith the contents of hydrogen and carbon dioxide to produce ammoniasynthesis gas at a high percentage hydrogen recovery without excessiveadsorption bed volume and with no or little external purge gas.

According to the invention a pressure swing adsorption process (PSA) forproducing ammonia synthesis gas from a raw gas containing hydrogen H₂,carbon dioxide CO₂ and at least one medium boiling point gas (MB)selected from the class consisting of nitrogen N₂, carbon monoxide CO,methane CH₄ and argon Ar, the nitrogen content being greater than thatrequired in ammonia synthesis gas, characterised by feeding to theadsorbent a raw gas in which hydrogen and total medium boiling pointgases are present in a volume ratio in the range 1.25 to 2.5, especially1.4 to 2.1 and the medium boiling point gases comprise nitrogen to theextent of at least 90% ^(v) /v on the total of such gases.

The raw gas fed to PSA may contain up to a few percent by volume ofwater vapour, depending on the capacity of the process to handle it, forexample as a result of including a silica gel water adsorption section.For the generality of PSA processes the water vapour content of the rawgas is under 1% ^(v) /v.

The MB components preferably comprises at least 95% ^(v) /v of N₂. Thehigh N₂ percentage and the substantial nitrogen affinity of theadsorbent have the effect that N₂ adequately displaces CO₂ duringregeneration of adsorbent.

Typical MB contents in the raw gas are, in % ^(v) /v on a dry basis:

    ______________________________________                                        CO             up to 2                                                        CH.sub.4       up to 5, especially up to 1                                    Ar             up to 1.                                                       ______________________________________                                    

(Other noble gases may be present. In practice helium and neon form partof the ammonia synthesis gas product and krypton and xenon part of theMB component, but their concentrations are too low to affect theoperation of the process).

The CO₂ content of the raw gas is preferably under 25% ^(v) /v on a drybasis. It may be substantially less as a result of a preliminary CO₂-removal treatment, but is preferably at least 10% as in raw gas not sotreated. For optimal operation of the process the volume ratio MB/CO₂ isin the range 1.3 to 2.5.

The composition of the raw gas is shown on the accompanying diagram,FIG. 1 hereinafter.

Corresponding to the high N₂ content of the MB components, thepercentage N₂ recovery is typically less than 80, for example in therange 45 to 65.

The PSA is broadly of the types described in U.S. Pat. Nos. 3430418,3564816 and 3986849, in which each bed takes part successively in thesesteps:

adsorption

co-current pressure equalisation (preferably multiple)

co-current partial depressurization to purge another bed

counter-current depressurization ("dump")

purge (optional)

re-pressurization.

At least 4 beds are used, preferably at least 5, in order to providemultiple pressure equalisations and thus increase product percentagerecovery. More preferably at least 10 beds are used.

In the ensuing definition of the special characteristics of PSA theterms "inlet" and "outlet" refer to the flow of gas during theadsorption step, and the terms "co-current" and "counter-current" meantoward such outlet or inlet respectively. When appropriate, beds areidentified by the index letters used in FIG. 3 accompanying, but this isfor ease of understanding and does not limit the invention to the cycleshown in that figure.

The PSA differs from those previously proposed in the compositions ofthe feed and product gases and in the detailed operations appropriatethereto. Preferably it includes the following features

(a) an adsorption step producing a product gas varying in compositionwith time. Typically the H₂ /N₂ ratio is low at the beginning becausethe bed (A) has been re-pressurized counter-current with an H₂ +N₂mixture and thus carries a relatively high loading of N₂ at its outletend. However, as adsorption flow proceeds this N₂ is gradually desorbedby gas from which components (CO₂, CH₄) other than H₂ have been adsorbedfurther upstream in the bed, and thus the H₂ /N₂ ratio increases. Withcontinuing adsorption flow, the N₂ -loaded part of the bed moves furthertowards the bed outlet and N₂ -breakthrough takes place. Such N₂-breakthrough is allowed to increase until the integrated H₂ /N₂ ratioover the whole adsorption step is at the level required in the productammonia synthesis gas. The adsorption flow is stopped well before theCO₂ adsorption front reaches the bed outlet. The variation of H₂ /N₂ratio with time is not necessarily symmetrical, but it is believedpreferable for the final ratio to be close (e.g. within 20%) of theinitial ratio. The highest ratio attained can be for example up to aboutdouble the lowest ratio. Consequently it is preferred to use a buffervessel for evening-out the product gas composition or (instead or inaddition) a PSA system having a plurality of adsorption beds inoperation simultaneously but out-of-phase. Using 3 such beds a 100%variation in ratio using a single bed can be decreased to 7%, typically.

(b) at the end of the adsorption step the bed contains these zones:

1. a zone containing mainly CO₂ as adsorbate;

2. a zone containing some CO₂ but partly loaded with MB;

3. a zone containing very little CO₂ but loaded with MB at a leveldecreasing from the zone boundary towards the bed outlet. The mainfunction of this zone is to remove N₂ incompletely from the gas so thatthe H₂ /N₂ integrates to the required level. If CH₄ is present in thegas leaving zone 2 it is removed in zone 3. At the end of the adsorptionstep zone 3 is still long enough to remain within the bed duringsubsequent co-current flow in pressure equalisation anddepressurization.

(c) at least one co-current downward pressure equalisation stepfollowing the end of an adsorption step in which the gas still in theadsorber as void space gas and adsorbed gas and initially at adsorptionpressure ("highest") is fed from the adsorber (A) outlet into the outletend of a fresh or previously regenerated bed (B). In the adsorber (A)this step advances the fronts of all those zones towards the outlet butzone 3 remains substantial in length. In the fresh or regenerated bed(B) this step moves the zones back towards the inlet, but also adsorbsN₂ in the outlet zone, since the N₂ partial pressure in the gas fed fromthe adsorber outlet is higher than in the purge gas (to be described)with which the fresh or regenerated bed was previously contacted.

The number of such co-current pressure equalisation steps is for exampleas follows:

(i) one, in a simple 4-bed system as depicted for example in FIG. 2 ofU.S. Pat. No. 3,430,418 and reproduced hereinafter;

(ii) two, in a modified 4-bed system as depicted for example in FIG. 2of U.S. Pat. No. 3,564,816;

(iii) two, in a 5-bed system as depicted for example in FIG. 3 of U.S.Pat. No. 3,430,418;

(iv) three, in systems using 6 or more beds, for example the 8-bed and10-bed systems described in U.S. Pat. No. 3,986,849.

In (ii) the two equalisation steps are separated by adepressurization-to-purge step (see below), but in (iii) and (iv) theyare consecutive.

The pressure after equalisation will be referred to as "intermediate",qualified by first, second etc in the event that more than oneequalisation step is used.

(d) a co-current depressurization-to-purge step following theequalisation step or steps, and providing a purge stream of gasthrottled from the intermediate pressure in the previouly equalisedadsorber (A), which is fed counter-currently through a bed (C) that hasjust been counter-currently depressurized (dumped). In the adsorber (A)this step further advances the main fronts of all three zones and also atail of each zone towards the outlet, but zone 3 still remainssubstantial in length, so that the gas leaving it is CO₂ -free. However,that gas is relatively rich in N₂, since H₂ was largely lost in theequalisation step or steps and N₂ is now desorbed as a result of lowerpressure and of displacement by advancing CO₂. In the dumped bed (C) thezones are moved further back towards the inlet and further CO₂ and MBgases are carried away, thus effecting a purge of this bed. The outletend of the dumped bed (C) adsorbs a significant quantity of N₂ from thepurge stream and, if the tail of the front of zone 1 has approached theoutlet of the dumped bed (C) in previous co-current depressurization,the purge stream now moves it backwards and decreases the risk of traceCO₂ breakthrough.

The flow of purge stream is stopped when the pressure in the adsorber(A) has fallen to a lower-intermediate level suitable for the operationof the step described in the next paragraph.

(e) a counter-current depressurization ("dump") step in which the gasremaining in the adsorber (A) at lower-intermediate pressure is releasedfrom the bed inlet. The resulting ("lowest") pressure is commonlyambient but can if desired be higher, or lower as a result of using avacuum pump: preferred pressures are described below. In the adsorberthe dump step moves the fronts of all three zones back towards the inletbut this effect is small towards the outlet end because there thequantity of gas flowing is small, and consequently any small "tail"quantity of CO₂ that approached the outlet end during pressureequalisation and co-current depressurization-to-purge tends to remainthere (see steps (d) and (f) and thus could be desorbed during asubsequent adsorption step into the product ammonia synthesis gas. Apartfrom such residual CO₂, the effect of the dump step is to expel asubstantial part of the CO₂ acquired in the adsorption step, and it ischaracteristic of the MB content of the raw gas that the MB gas presentin the bed before the dump step is in balance with what is required todisplace such CO₂. As shown in accompanying FIG. 2, an H₂ /MB ratiolower than the specified lower limit would exact a severe penalty in thebed volume needed, and a ratio higher than the specified upper limitwould lead to inadequate removal of CO₂. The fraction of the adsorbedCO₂ that is expelled in the dump step is greater, the lower the H₂ /MBratio. The remainder of the adsorbed CO₂ is removed in one or more purgesteps.

(f) a counter-current purge step using gas from a bed (D) undergoingco-current depressurization. The effect of this purge on the dumped bed(C) was described in paragraph (d) above. If the H₂ /MB molar ratio inthe raw gas is in the lower part of the defined range, this purge may beunnecessary or may be adequately effected using only a part of the gasavailable from the co-current depressurization; if desired, co-currentdepressurization gas, dump gas, initial purge gas and final purge gascan be fed out to different uses appropriate to their compositions.Usually it is preferred to use this purge step to ensure desorption ofthe tail of the CO₂ front, to help establish a quantity of adsorbed N₂at the end outlet and to keep the CO₂ main front well back from the bedoutlet in the event that it has advanced as a result of chance raw gascomposition fluctuations; the latter is of concern mainly at higher H₂/MB ratios in the defined range.

(g) a counter-current purge step using gas from outside the PSA cycle.Such "external" purge is usually unnecessary and is envisaged as aremedy in the event of excessive advance of the CO₂ front duringadsorption. Gases conveniently usable are ammonia synthesis loop purgegas or adsorber product gas or some gas external to the whole ammoniaproduction plant. However, it is a major advantage of the invention thatthe nitrogen purge specified in GB-A-2103199 and BE-A-885126, whichnecessitates an air-separation plant and makes air-reforming ofhydrocarbon feedstock unsuitable, is not used.

If in (f) or (g) a remedial purge is to be operated, this can beintermittent and a spare bed can be provided.

(h) an upward counter-current pressure equalisation step in which thepurged adsorber (A) receives at its outlet end the gas from an adsorber(B) that has ended its adsorption step, as described in paragraph (c)above. As a result of the direction in which the gas enters the bed, thefronts in it are moved back towards its inlet, and the zone nearest itsoutlet is brought towards equilibrium with the relatively N₂ -rich gasdelivered by the downward-equalising adsorber.

(i) a counter-current re-pressurization step in which product gas isdiverted from the outlet stream into the outlet of the adsorber (A)during or after upward pressure equalisation. The feed of product gastakes place preferably throughout the eqasation as well as after it, sothat the rate of flow of such gas does not vary much during the wholecycle. Alternatively or additionally the product gas feed may be from areservoir. At the end of re-pressurization the adsorber is ready toreturn to adsorption duty.

(j) the adsorption ("highest") pressure is in the range 25 to 50,especially 30 to 40, bar abs. and the purge and final dump ("lowest")pressure is over 1, especially in the range 3 to 5 bar abs. The ratio ofhighest pressure to lowest pressure is preferably in the range 8 to 25,for example 10 to 15.

Further processing of the purged and dumped gas is described below.

The adsorbent used in the PSA can be chosen from among availablematerials including varieties of active carbon, zeolites and silica gel,in respect of which gas adsorption data are published or are availablefrom companies specialising in adsorption. Among the zeolites, those ofpore diameter 5 Angstrom units and less are generally useful in view ofthe small size of the molecules involved, for example calcium zeolite A.Molecular sieves providing substantially increased adsorption of COrelative to N₂, for example by a factor of more than 10, such asmordenite appear despite their larger pore diameter, to be potentiallyof value but are not normally needed.

Whereas the CO content of the raw gas is specified as less than 2% ^(v)/v on a dry basis, it is preferably substantially less, especially under0.5% ^(v) /v. Thus the shift stage producing the raw gas preferablyincludes low temperature shift over a copper-containing catalyst atunder 250° C. outlet temperature. To ensure a synthesis gas CO contentlow enough to avoid poisoning an iron catalyst used in ammonia synthesisthe product gas from the PSA system is preferably methanated, also asdescribed hereinafter. If a ruthenium ammonia synthesis catalyst is tobe used, the CO content is less critical and such post-PSA methanationmay be unnecessary. To ensure a very low raw gas CO content, for exampleunder 0.01%, the raw gas is subjected preferably to selective catalyticoxidation or methanation as described hereinafter.

The invention provides also a combination process comprising

(a) reacting a carbonaceous feedstock with steam and an O₂ -N₂ mixtureusing process conditions and reactant proportions so as to produce acrude gas in which the ratio by moles ##EQU1## is in the range 1.25 to2.5, especially 1.4 to 2.1 and the MB gas consists of N₂ to the extentof at least 90% ^(v) /v;

(b) subjecting the crude gas to catalytic shift reaction with steam toconvert CO substantially to CO₂ +H₂ ; and

(c) removing CO₂ and MB by PSA as hereinbefore defined, whereby toproduce a N₂ /H₂ ammonia synthesis gas.

In the crude gas leaving stage (a) the CO+CO₂ content is preferably inthe range 10-25% ^(v) /v on a dry basis and the MB gases are preferablyN₂ to the extent of at least 90%.

In step (b) the conversion of CO is such as to have less than 2,especially less than 0.5, % ^(v) /v of CO on a dry basis.

At some stage in the combination process before PSA provision is made toremove any sulphur compounds that may be introduced with the feedstock.Sulphur compounds removal can be applied to crude gas or shifted gas butin preferred processes in which stage (a) involves catalytic reaction ofa volatilisable feedstock, is applied to the feedstock before stage (a).

Stage (a) could in principle be carried out in a single reaction of allthree reactants but when catalysed is preferably carried out in twoparts, one involving feedstock and steam, the other involving also O₂.

In one form of this stage, a volatile hydrocarbon feedstock is reactedwith steam over a catalyst heated externally ("primary reforming") toproduce a gas containing CO, CO₂, H₂ and CH₄ and the resulting gas isreacted with the O₂ -N₂ mixture adiabatically to convert CH₄ to CO+H₂and introduce N₂ ("secondary reforming"). Such a sequence resemblescrude synthesis gas generation as described in our U.S. application No.4,298,588; if desired, the reaction with steam could be carried out bypreheating followed by adiabatic reaction, as described in our U.S.application No. 4,303,982.

In a preferred form of stage (a) the heat required for primary reformingis obtained by indirect heat exchange with the hot gas resulting fromsecondary reforming. It is fortunate that the heat balance of the tworeforming stages is such that when using air or the O₂ -N₂ mixture the(CO+H₂)/(MB-CO) ratio and N₂ content in the specified range can bereadily attained. However, the invention includes also the use ofmoderately enriched air containing up to 35% of O₂ and O₂ -depleted aircontaining down to 15% v/v O₂. Thermodynamic data permitting calculationof temperatures, pressures and reactant proportions are readilyavailable to chemical engineers.

Upstream of PSA there can be partial removal of CO₂, especially when theH/C atomic ratio is less than 3 in the hydrocarbon feedstock.

In a further form of stage (a) the feedstock is methanol and is reactedwith steam and air in a single catalytic operation.

After PSA and after or before any final CO methanation the PSA productgas is compressed to ammonia synthesis pressure, which is usually over40 bar abs, for example in the range 40-120 bar abs as in recentlyproposed processes or in the range 120-250 bar abs as in most processesin industrial operation at present. Especially when a centrifugalcompressor is used, it is important to avoid gross variations in themolecular weight of the synthesis gas, and therefore the PSA system mayinclude a buffer vessel or 3 or more simultaneous adsorbers out of phaseas already described. For the sake of a steady flow rate to thecompressor the re-pressurization of the adsorber before the adsorptionstep is by a steady slow flow of PSA product gas.

The ammonia synthesis stage can be of any convenient type and may,indeed, be carried out in existing plant designed and used inconjunction with conventional synthesis gas generation. It willtypically be subject to one or more of the following detailedmodifications:

(a) unless a product gas H₂ /N₂ ratio different from 3 has been chosen,the synthesis purge gas rate will be very small and will not justifypurge gas separation for recovery of ammonia and H₂. Conveniently suchpurge gas is recycled to synthesis gas generation;

(b) since the PSA product gas is very pure and non-reactants do notaccumulate in the circulating synthesis gas, the partial pressures of N₂and H₂ are higher and thus the rate of production of ammonia is higher.Alternatively energy can be saved by decreasing the gas circulationrate.

Other combinations designed for energy recovery are described below.

The invention provides a further combination process in which the CO₂-rich waste gas stream is concentrated and fed to one or more of thefollowing processing stages:

(a) production of solid CO₂ ;

(b) production of liquid CO₂ ;

(c) production of urea.

Such concentration of CO₂ can be by means of an adsorptive treatment orwet treatment, for example by adsorption in a regenerable liquid such asaqueous potassium carbonate optionally containing an activator, analkanolamine or a pressure-sensitive solvent. If the CO₂ is to be usedfor urea production it can be recovered as ammonium carbonate. If theCO₂ -rich stream is dumped and purged at the preferred pressure 3-5 barabs. it can be subjected to concentration without compression; however,the invention includes a compression stage if the dump and purgepressure is lower or if later processing requires higher pressure.Although the PSA system would be less expensive if CO₂ were removedupstream therof, the disadvantages would be incurred that the wholeammonia production process would include a wet stage; the necessary CO₂removal plant would have to handle much larger gas volumes at higherpressure, and an additional pressure-drop would be incurred before PSA.

A result of concentrating the CO₂ is to produce a stream containing H₂and MB gases which is a much more convenient fuel than the CO₂ -richstream and is more uniform in composition.

Further combination processes designed to afford energy economy aredescribed hereinafter.

The catalytic shift reaction should preferably be of the "clean" type,e.g. when selective oxidation or methanation is to precede PSA. Theshift reaction can be carried out in conventional ways, for example

"high temperature", with an inlet temperature of 330°-400° C., outlettemperature 400°-500° C., usually over an iron oxide/chromia catalyst,and affording an outlet CO content in the range 2-4% ^(v) /v on a drybasis so that a further stage is needed;

"low temperature", with an inlet temperature of 190°-230° C., outlettemperature 250°-300° C., usually over a catalyst comprising metalliccopper, zinc oxide and one or more other difficultly reducible oxidessuch as alumina or chromia, and affording an outlet CO content in therange 0.1 to 1.0 especially under 0.5% ^(v) /v on a dry basis;

"combination", using the sequence of high temperature shift, cooling byindirect heat exchange and low temperature shift; if desired, eithershift stage can be subdivided with interbed cooling.

Alternatively a "medium temperature" shift can be used, in which theinlet temperature is in the range 250°-325° C. and the outlettemperature up to 400° C. A suitably formulated supported coppercatalyst can be used. The outlet CO content is up to 2.0% ^(v) /v on adry basis.

Whichever shift method is used, it is preferably operated in indirectheat exchange with a coolant, especially water under pressure. Thus thecatalyst can be disposed in tubes surrounded by the water, or viceversa. Utilisation of the heat taken up by the water may be bygenerating steam at for example 15-50 bar abs. pressure and use of suchsteam as feed to the shift stage or in generating the CO containing gasfed to shift.

Preferably the shift stage in heat exchange with water is characterisedby controlling the water flow rate so that incomplete vaporisation takesplace, and contacting the resulting steam/water mixture with a gaseoushydrocarbon, whereby to form a mixture thereof with water vapour. Analternative shift stage is carried out by

(a) reacting carbon monoxide with steam over a catalyst in indirect heatexchange with boiling water under pressure in a first circuit;

(b) condensing the resulting steam in indirect heat exchange with waterin a second circuit and returning the resulting first circuit condensateto the indirect heat exchange in stage (a);

(c) contacting the resulting optionally partly boiling hot secondcircuit water with a gaseous hydrocarbon, whereby to form a mixturethereof with water vapour.

More particularly the gaseous hydrocarbon/water vapour mixture is fed toa gasifier, especially a catalytic endothermic steam reformer and/orcatalytic partial oxidation, and the product of the gasifier undergoesthe shift reaction in heat exchange with the water. The shifted gas iscooled further, preferably by direct heat exchange with water. In theshift stage first mentioned water returned from the contacting andmake-up water are conveniently fed to the water side of the heatexchange surfaces in the reactor. In the alternative shift step returnedwater and make-up water become the second circuit water. To maintain thethermal balance of the process and to adjust the temperatures of thegases to be contacted with water, the hot steam/water mixture or secondcircuit water may be indirectly heat exchanged with the product of thegasifier before it enters the shift steps; warm water may be produced bydirect cooling of shifted gas and may be indirectly heat exchanged withshifted gas.

The shifted gas is finally cooled to below the dewpoint of steam, liquidwater is separated from it and stages are carried out to remove bulk andresidual CO₂ and residual CO from it: preferably PSA is used butconventional steps of CO₂ removal in a regenerable liquid absorbent andfinal carbon oxides removal by catalytic methanation can be used, in amodified process.

To provide power requirements of process sequences producing gas at highpressure the invention provides a combination process having the stagesof

(a) reacting a hydrocarbon feedstock with steam over an externallyheated catalyst;

(b) subjecting the product of stage (a) to a shift reaction;

(c) removing carbon oxides from the shifted gas;

(d) recovering heat from the hot product gases of stage (a) and/or stage(c) and/or from combustion gases formed in providing the externalheating in stage (a), and using such recovered heat to power an enginedriving a compressor for the product of stage (c) and/or for a gastaking part in stage (a): characterised by

(i) in stage (d) recovering said heat by heat exchange with reactantsand/or liquid water but substantially without producing external steam;

(ii) carrying out stage (c) by selective adsorption, producing a CO₂-containing off gas including combustibles and using it to power a gasexpander producing at least part of the compression power in stage (d).

In stage (i) the process can produce a steam-water mixture (to be usedin a saturator), which is to be distinguished from external steam fed toa turbine or mixed in gaseous form with the feedstock in stage (a) orused in some way outside the process stages listed. Alternatively steamis raised only in a closed ("first") circuit and used to heat the waterto be used in the saturator.

The energy recovery in stage (ii), including heat recovery from gasturbine effluent, usually needs to be supplemented by energyimportation, such as in an independently fuelled gas turbine or ofelectricity, but the over-all energy consumption is typically less thanin a conventional process with heat recovery by steam raising.

After PSA there may be further purification steps, depending on whetherthe ammonia synthesis catalyst is of the ruthenium or the iron type.Thus:

1. For a ruthenium catalyst the sub-residual CO is unlikely to be apoison, hence no further purification is needed. However, the CO canreact to methane over the ammonia synthesis catalyst, thus possiblyproducing an undesired exotherm;

2. If the gas from PSA is methanated before the ammonia synthesis thedisadvantage of an undesired exotherm in the synthesis is avoided, andthe water produced in methanation does not affect the rutheniumcatalyst;

3. If an iron synthesis catalyst is to be used, the gas from PSA may bemethanated and dried before it reaches that catalyst. Drying can be bymeans of a regenerable solid adsorbent, but most conveniently the gas iscontacted with liquefied ammonia, suitably as described below

If post-PSA methanation is used, the "sub-residual" CO content of thegas leaving PSA is typically in the range 0.05 to 0.20% ^(v) /v.Methanation is carried out typically at 250°-400° C. over a supportednickel or cobalt or possibly ruthenium catalyst.

Methanation takes place conveniently after any compression of the gasfrom the synthesis gas generation pressure (10-60 especially 25-50 barabs.) to ammonia synthesis pressure (30-300, especially 40-120 barabs.), because the gas leaving PSA is very dry.

The methanated gas contains 0.05 to 0.2 ppm ^(v) /v of water vapour andis free of CO₂. Consequently it can be adsorbed by liquid ammoniawithout significantly contaminating it. For this purpose the methanatedgas is mixed with reacted synthesis gas before, or after partial,removal of ammonia therefrom, and the mixture is cooled and passed to anammonia separation catchpot.

If the CO content of the gas entering PSA is decreased by selectiveoxidation, the oxidising agent can be air, enriched air or highconcentration oxygen, depending on how much nitrogen can be accepted inthe PSA feed gas. The selective oxidation catalyst is suitably supportedplatinum (0.01 to 2.0% ^(w) /w) containing possibly one or more ofmanganese, iron, cobalt or nickel as a promoter. A description of asuitable selective oxidation process is given in UK 1555826 and in thearticles by Colby et al (23rd Symposium on safety in ammonia plants andrelated facilities, Am. Inst. Chem. Engrs. Conv., Miami, Nov. 1978) andBonacci et al. (Am. Inst. Chem. Engrs. Symposium, Denver Aug. 1977). Theinlet temperature is typically below 50° C. and the outlet temperatureunder 100° C., especially under 80° C., in order to assure highselectivity for CO oxidation and against H₂ oxidation.

To ensure thorough CO removal, possibly making post-PSA methanationunnecessary, the selective oxidation is preferably carried out in heatexchange with a coolant or, more conveniently in a plurality of stagesin succession, each fed with oxidant sufficient to oxidise part of theCO initially present but insufficient to produce an adiabatictemperature rise to over 80° C., and with inter-stage indirect heatexchange cooling.

If the CO content of the gas entering PSA is decreased by selectivemethanation the shifted gas is passed to the methanation stagepreferably without removing unreacted steam, and conveniently without achange in temperature. This makes possible a valuable simplification, inthat the methanation catalyst can be disposed in an adiabatic bed at thedownstream end of the reactor in which the cooled shift stage or thelowest temperature shift takes place.

The combination shift+methanation reactor consitutes a further featureof the invention.

The conditions favouring methanation of CO preferentially to CO₂ includethe following

relatively low temperature, for example under 300°, especially under200° C.;

relatively high pressure, for example over 20, especially over 30 bar.

noble metal, especially ruthenium-containing, catalyst.

In order to control temperature, the catalyst may be disposed in a zonehaving indirect heat exchange surfaces, for example in tubes surroundedby coolant or in a bed having tubes through which coolant is circulated;however, at the preferred shift outlet CO content an adiabatic bedsuffices.

After methanation, the gas is cooled to remove unreacted steam ascondensate, then passed to PSA.

The ammonia synthesis system in which the gas is used involves synthesisat a pressure typically in the range 30-300, especially 40-120 bar abs.,and a catalyst outlet temperature preferably in the range 300°-450° C.The catalyst can be metallic iron with one or more promoter oxides suchas those of potassium and aluminium, and possibly containing up to 10%of cobalt (calculated as Co₃ O₄ on a composition in which iron iscalculated as Fe₃ O₄). Alternatively the catalyst can be a supportedplatinum group metal, for example ruthenium supported on graphite. TheH₂ /N₂ ratio of the gas entering the synthesis catalyst is preferably inthe range 2.7 to 3.0 when an iron catalyst is used but can be lower, forexample down to 2.2 if provision is made to recycle part of theunreacted gas as purge to synthesis gas generation, or to recoverhydrogen from such purge gas. Such ratios are suitable also for aruthenium catalyst, but still lower ratios have been proposed formulti-stage ammonia synthesis. Usually the synthesis includes cooling ofreacted synthesis gas, separation of liquid ammonia and recycle ofunreacted gas to the synthesis catalyst, but the multi-stage synthesiscan be on a once through basis.

The invention is illustrated by the accompanying drawing in which

FIG. 1 is a triangular gas composition diagram showing preferred raw gascompositions to be fed to the PSA system;

FIG. 2 is a diagram indicating the variation of adsorber bed volume andpercentage CO₂ retention as a function of H₂ /MB ratio;

FIG. 3 shows the steps of an illustrative PSA system usable in theprocess of the invention;

FIG. 4 shows a combination process including raw gas generation, shiftwith heat recovery. PSA, methanation and ammonia synthesis;

FIG. 5 shows a combination process with selective methanation of CObefore PSA; and

FIG. 6 shows a combination process with selective CO oxidation beforePSA.

In FIG. 1 the quadrilateral ABCD represents the H₂ /MB ratio range 1.25to 2.5 in combination with the preferred CO₂ content range 10-25% ^(v)/v, and quadrilateral EFGH the particularly preferred H₂ /MB range 1.5to 2.1 in combination with the MB/CO₂ range 1.3 to 2.5. In each case theMB gas is at least 90, especially at least 95% ^(v) /v N₂.

In FIG. 2 the curve relating to the left-hand axis represents ourdiscovery that there is a region II represented by our defined H₂ /MBratio range over which the bed volume is markedly less sensitive to H₂/MB ratio. This is highly valuable in that a particular PSA plant designis usable for the variety of feed gas compositions that may result fromchoice of feed gas generation plant and feedstock and unavoidablefluctuations in gas composition. At lower H₂ /MB ratios the bed volumeis greater (region III), which results in greater expense and decreasedhydrogen recovery, and is also much more sensitive to feed gascomposition fluctuations. In region I the bed volume is less, but (seeright-hand axis) the CO₂ retained in the gas is excessive; this problemhas previously been solved (in proposals) by the expensive expedient ofan external purge.

In FIG. 3 the inlet and outlet streams are numbered as in the flowsheetsof FIGS. 4-6 accompanying. The above definition of the characteristicsof the PSA system serves as a description of this figure.

Table 1 shows temperatures, pressures, gas compositions and flow ratesfor the PSA unit, and uses the same numerals.

The product stream "56 integrated" is the resultant of a stream in whichfor each individual adsorption step the H₂ /N₂ ratio at an intermediatetime was about twice as high as initially and finally; in a 10 bedsystem using 3 adsorbers out of phase, this variation range is onlyabout 7%.

As a result of the residual CO and CO₂ content of stream 56 integratedit is suitable for ammonia synthesis over a ruthenium catalyst, butshould be methanated, possibly after compression, before synthesis overon iron catalyst.

                                      TABLE 1                                     __________________________________________________________________________           Temp.   Gas composition % v/v                                                                           .sup.+ Flow                                  Position                                                                             °C.                                                                        *Press.                                                                           CO CO.sub.2                                                                         H.sub.2                                                                          CH.sub.4                                                                         N.sub.2                                                                          Ar rate                                         __________________________________________________________________________    54 inlet                                                                             35  35  0.31                                                                             16.07                                                                            53.09                                                                            0.48                                                                             29.43                                                                            0.61                                                                             7998                                         56 integrated                                                                        40  34  0.08                                                                             0.01                                                                             74.46                                                                            0.06                                                                             24.82                                                                            0.58                                                                             5493                                         57     30  1.5 0.82                                                                             51.29                                                                            6.27                                                                             1.38                                                                             39.55                                                                            0.69                                                                             2505                                         __________________________________________________________________________     *bar abs.                                                                     .sup.+ kg mol h.sup.-1                                                   

In the process shown in FIG. 4 desulphurised natural gas 10 is at 11mixed with a synthesis purge stream (to be described) and fed at 12 tothe upper (saturator) section of tower 14. It contacts a hot waterstream (to be described) fed in at 16 over the upper packed saturationzone, then is mixed with more steam fed in at 18. The resulting warmsteam/gas mixture (S/G ratio 2 to 5, temperature 150°-250° C., 25-50 barabs) is preheated at 20 (350°-550° C.) and fed at 22 into annular bedsof steam reforming catalyst (supported Ni or Co) heated externally infurnace 24. The resulting hot gas (600°-800° C.) containing CO, CO₂, H₂,unreacted steam and several percent of CH₄ passes into the closed end 26of tube 22 and returns through inner tube 28 in which it gives up heatto the reacting gas in the annular catalyst bed. The resulting partlycooled gas (450°-650° C.) is fed at 30 into the uppermost (combustion)section of furnace 24, where air or O₂ -enriched air 32 is fed into it.A flame is formed and the combustion products are brought to equilibriumat a lower CH₄ content over secondary reforming catalyst 36. Theresulting gas, still at 900°-1050° C., is the source of heat for theouter tubes containing the annular beds of steam reforming catalyst. Inheat exchange with these tubes the gas is cooled typically to 450°-65°C. It leaves furnace 24 at 38, is cooled in heat exchanger 20 and waterheater 39 to shift inlet temperature and passes into water-cooled shiftreactor 40 in which the catalyst is disposed in tubes surrounded bywater in a pressure shell to be described. In reactor 40 the shiftreaction is brought substantially to equilibrium at a temperaturecontrolled at typically 230°-280° C. and an outlet CO content in therange 0.1 to 1.0% ^(v) /v on a dry basis. The outlet temperature ispreferably lower by 10°-30° C. than the inlet temperature. The shiftedgas is cooled at 41, passed into the lower packed desaturation zone oftower 42 and therein contacts cool water fed in at 44 from a source tobe described. The resulting waterdepleted gas is fed out, and cooled at50 to below the dewpoint of steam. Optionally it may, before suchcooling, be passed into reactor 48, in which CO is oxidised selectivelyas described below with reference to FIG. 6. From the gas at below thedewpoint of steam, liquid water is separated in catchpot 52, from whichdry gas is taken overhead and passed into selective adsorber 54.Adsorber 54 includes essentially a bed of material such as active carbonor molecular sieve on which CO₂ is strongly adsorbed, H₂ is very weaklyif at all adsorbed and MB gases are partly adsorbed. It also includessuch a bed under regeneration and possibly other beds undergoingtreatments such as temperature and pressure adjustments, and thenecessary change-over valves. From 54 purified ammonia synthesis gas ispassed out at 56 to a synthesis gas compressor 80 and thence to anammonia synthesis plant to be described. A regeneration off gas,containing CO₂, N₂, some H₂ and possibly CO and CH₄ is passed out viapoint 57, whereafter it receives a feed of combustible such as naturalgas and possibly synthesis purge gas, to the compressor and combustor(not shown) of gas turbine 58 providing the shaft power for aircompressor 59. The composition of the regeneration off-gas will dependon the extent to which the oxidant gas fed at 34 has been O₂ -enrichedbefore entering compressor 59. If there has been no enrichment, it willbe rich in N₂ and possibly also in CH₄, since the rate of feed of air at32 may be kept below that necessary to react all the CH₄, in order toavoid excessive N₂. Further, although it is simplest if the gas passedout at 56 is stoichiometric for ammonia synthesis, it may be preferredto pass out at 56 a gas having an H₂ :N₂ ratio under 3, and to removethe excess N₂ and any CH₄ cryogenically before or after compressing thegas. Alternatively the excess N₂ and any CH₄ can be removed afterammonia synthesis.

The ammonia synthesis gas passed out at 56 is compressed to ammoniasynthesis pressure at 80 and (possibly via a feed/effluent heatexchanger not shown) fed into methanator 82 in which its sub-residualcontent of CO is converted to CH₄. The methanated gas is cooled at 84°to 30°-40° C. and mixed with ammonia-containing reacted synthesis gas at86. The mixture is at 88 chilled to below the dewpoint of ammonia andpassed into catchpot 90, in which liquid ammonia is separated. Theunreacted gas overhead is divided into a main recycle stream and a purgestream. The main recycle stream is compressed by a few bars incirculator 94 and passed into synthesis reactor 96 which includes one ormore catalysts beds (Fe or Ru), quench or indirect temperature controlmeans, a feed/effluent heat exchanger and external heat recovery meanssuch as a boiler, superheater or water heater. The gas after reaction isfed to point 86 already described. The purge stream is fed to point 11,where it joins the natural gas feedstock to the process. (If CH₄ ispresent it is converted to raw synthesis gas in reformer 24. If N₂ ornoble gases are present, they pass through the reformer and shiftreactor but are adequately removed adsorptively at 54. Although dilutionof the reformer feed by such non-reacting gases takes place, it has theadvantage of increasing the steam-to-CH₄ ratio obtainable by saturationin tower 14 and, thus decreasing or removing the need for steam feed 18.Further, such recycle to point 11 means that adsorber 54 is the onlymeans necessary for removing non-condensible gases from the process).

The water system of the process receives cold condensate as the bottomsof catchpot 52 and cold make-up water 63 at 64 and feeds the mixture viapump 66 to point 68 where it is united with a cool stream to bedescribed and whence the whole mixture is fed at 44 over the lower(desaturator) packed section of tower 14. Here the unreacted steam inthe shifted gas from reactor 40 condenses into the water, giving a warmwater stream which is taken as bottoms and fed via pump 70 to 3 heatingstages, by indirect heat exchange first with shifted gas at 41, then at43 with condensing steam raised in shift reactor 40, then with partlycooled secondary reformer gas at 39. It may then be still entirelyliquid or may be partly boiling and is fed at point 16 to the saturationzone of the upper section of tower 14. The cooled water remaining aftercontacting in the upper section of tower 14 is cooled at 67 by heatexchange with boiler feed water to be treated in a de-aerator (notshown), then united with cold condensate at point 68.

Table 2 shows temperatures, pressures and gas flow rates in a processfor producing 1100 metric tons of ammonia per day.

                                      TABLE 2                                     __________________________________________________________________________            Temp Pressure                                                                           Gas flow rate kg mol h.sup.-1                               Position                                                                              °C.                                                                         bar abs.                                                                           CO  CO.sub.2                                                                          H.sub.2                                                                            CH.sub.4                                                                          N.sub.2                                                                           Ar  H.sub.2 O                                                                         NH.sub.3                       __________________________________________________________________________    *10     30   46   --  5.14                                                                              37.02                                                                              1367.37                                                                           47.14   --  --                             +(11    30   101  --  --  298.89                                                                             18.65                                                                             99.63                                                                             28.18                                                                             --  18.99)                         22 inlet                                                                              425  40   --  5.14                                                                              37.02                                                                              1367.37                                                                           47.14   3712                                                                              --                              32     670  39   --  --  --   --  2453.13 31.37                                                                             --                             38      540  37   972.15                                                                            492.94                                                                            3537.43                                                                            23.76                                                                             2500.27 3100.61                                                                           --                             40 outlet                                                                             230  36   27.47                                                                             1437.62                                                                           4482.11                                                                            23.76                                                                             2500.27 2155.93                                                                           --                             +52                                                                              overhead                                                                           35   35.3 27.47                                                                             1437.62                                                                           4781.00                                                                            42.41                                                                             2568.84                                                                           59.24                                                                             15.13                                                                             --                             56      35   35   8.93                                                                              --  4436.37                                                                            11.88                                                                             1423.10                                                                           32.69                                                                             --  --                             57      35   1.5  18.54                                                                             1437.62                                                                           344.63                                                                             30.53                                                                             1145.74                                                                           36.55                                                                             15.13                                                                             --                             82 outlet                                                                             180  130  --  --  4409.58                                                                            20.81                                                                             1469.86                                                                           32.69                                                                             8.93                                                                              --                             96 inlet                                                                              260  107  --  --  18658.80                                                                           1036.20                                                                           6201.25                                                                           1734.39                                                                           --  1055.60                        96 outlet                                                                             432  104  --  --  14952.56                                                                           1036.20                                                                           4845.83                                                                           1734.39                                                                           --  3766.43                        __________________________________________________________________________     *This starting gas includes H.sub.2 added before desulphurisation and als     53.93 kg mol h.sup.-1 of C.sub.2 + hydrocarbons additional to the CH.sub.     shown.                                                                        +Whereas synthesis purge gas (composition as 96 inlet) could with             advantage be added at 11, the calculated flow rates are based on addition     point 53.                                                                      This is process air including 652.5 kg mol h.sup.-1 of O.sub.2.         

In the process shown in FIG. 5 desulphurised natural gas 10 mixed at 11with ammonia synthesis purge gas is fed at 12 to the bottom of saturatortower 14. It contacts a hot water stream (to be described) fed in at 16over the packed saturation zone, then may be mixed with more steam fedin at 18. The resulting warm steam/gas mixture is reacted and cooled toshift outlet temperature by the stream described with reference to FIG.4 and passed into water-cooled shift/methanation reactor 40 in the upperpart of which the shift reaction takes place. The resulting gas passes,via an optional cooler or water injection (not shown) into supportedruthenium catalyst 48 in the lower part of reactor 40, in which CO isreacted with H₂ to produce CH₄ but CO₂ remains unreacted. The methanatedgas is cooled at 49 in heat exchange with water to be described, thancooled at 50 to below dewpoint of steam, whereafter liquid water isseparated in catchpot 52. (If desired the cooled gas from 49 can becontacted with cool water before 50). Dry gas taken overhead from 52 ispassed into selective adsorber 54, as in FIG. 4. The gas enteringadsorber 54 may include moist synthesis purge gas-H₂, N₂, CH₄ -as fed tocatchpot 52 at 62 instead of or in addition to point 11. The compositionof the regeneration off-gas will depend on the factors described withreference to FIG. 4.

The water system of the process receives cold condensate as the bottomsof catchpot 52 and cold make-up water 63 at 64 and feeds the mixture viapump 66 to point 68 where it is united with a cool stream from saturator14. The whole mixture passes to pump 70 and thereafter is heated in 3steps, by indirect heat exchange first with methanated gas at 49, thenwith condensing steam at 43, then with partly cooled secondary reformergas at 39. It may then be still entirely liquid or may be partly boilingand is fed at point 16 to saturator tower 14.

In a typical process according to FIG. 5 a shift outlet gas of % ^(v) /vcomposition CH₄ 0.22, CO 0.26, 13.5, H₂ 42.0, N₂ 23.5, H₂ O 2.4 at 230°C., 36 bar abs. pressure, is fed to a supported ruthenium methanationcatalyst. The outlet CO content is to be of the order of 25 ppm ^(v) /vand the CO₂ is to be methanated to the extent of at most one hundredthpart. The resulting gas is suitable for conversion to ammonia synthesisgas by drying, then selective adsorption of CO₂ completely, CH₄ partly,N₂ to the extent of one-third and H₂ to the extent of one-fiftieth part.The ammonia synthesis gas then needs no further purification and is fedto an ammonia synthesis loop.

In the process according to FIG. 6 desulphurised natural gas 10 is fedat 12 to the upper (saturator) section of tower 14. It contacts a hotwater stream (to be described) fed in at 16 over the lower packedsaturation zone, then is both heated and mixed with more steam bycontact with a steam/water mixture (to be described) fed in at 19. The2-stage saturator may make steam addition at 18 unnecessary. Theresulting warm steam/gas mixture is treated as in FIG. 4 down to thewater removal stage. The waterdepleted gas is fed out at 46, cooledfurther, typically to 30°-50° C. (by means not shown), and passed intoreactor 48, in which CO is oxidised catalytically by air or enrichedair. The gas may be cooled to oxidation catalyst inlet temperature andsubjected to a further oxidation stage (not shown). It is cooled at 50to below the dewpoint of stream, whereafter liquid water is separated incatchpot 52, from which dry gas is taken overhead and passed intoselective adsorber 54, as in FIG. 4. Adsorber 54 can receive at 60 afeed of N₂ separated from ammonia synthesis reacted gas. If O₂ -enrichedair if fed at 34, this is the product of air separation, which would beupstream of compressor 59 and would include a pre-compressor driven byturbine 58. It would give also an N₂ stream suitable for feeding atpoint 60. Turbine 58 exhausts via heat recoveries (not shown).

The water system of the process receives cold condensate as the bottomsof catchpot 52 and cold make-up water 63 at 64 and feeds the mixture viapump 66 to point 68 where it is united with a cool stream to bedescribed and whence the whole mixture is fed at 44 over the lower(desaturator) packed section of tower 14. Here the unreacted steam inthe shifted gas from reactor 40 condenses into the water, giving a warmwater stream which is taken as bottoms and sent by pump 70 to point 72.From point 72 a first stream of warm water is fed to lower packedsaturation zone at 16 and a second stream is fed into the shell ofreactor 40, in which it is brought just to the boil but without forminga distinct steam phase. The steam/water mixture emerging from the shellof reactor 40 is fed to the upper packed saturator section of tower 14,where it partly evaporates and completes the addition of steam to thenatural gas feed. Residual hot water is taken as bottoms from the uppersection of tower 14 and cooled at 76 in heat exchange in a closedcooling circuit (not shown) or with boiler feed water fed in at 78,after being warmed at 76, deaerated at 80 and sent out at 82 to furtherheat recoveries in the effluent of gas turbine 58. The cooled saturatorwater from 76 is united at 68 with cold condensate and make-up water. Asan alternative, heat exchanger 76 could be a cooling coil in the bottomof the upper section (saturator) of tower 14.

The water/steam system of FIG. 6 could be used in the processes of FIGS.4 and 5, and vice versa. When cooling is by steam raising in a closed("first") circuit filled with high quality boiler feed water; the tubesand shell of reactor 40 can be made of carbon steel, but when the watercirculated through the shift reactor shell in the process of FIG. 6 alsopasses through tower 14 and contains CO₂ the tubes and shell arepreferably made of corrosion-resistant alloy.

In a further alternative form of tower 14 (dotted path), the heatexchange in shift reactor 40 generates a steam/water mixture, and thusis fed to the upper bed of the saturator section of the tower. However,between the two beds of the saturator section there is a chimney-platefrom which water passes back to point 73, where it receives make-up warmwater from point 72 and then re-enters the shift reactor water shell.Natural gas partly saturated by warm water fed at 16 as before passes upthrough the chimney plate and is finally, heated and saturated in theupper bed.

We claim:
 1. A process for the production of ammonia comprising(A)forming ammonia synthesis gas by(a) reacting a carbonaceous feedstockwith steam and a gas containing oxygen and nitrogen using processconditions and reactant proportions so as to produce a crude gascontainingunreacted steam; hydrogen; medium boiling gas consistingofcarbon monoxide; nitrogen in an excess of that required in the ammoniasynthesis gas;methane; and, optionally, argon; and in which the molarratio of the sum of hydrogen and carbon monoxide to nitrogen, methane,argon, if any, is in the range 1.25 to 2.5, and the content of carbonmonoxide plus carbon dioxide, if any, is in the range 10 to 25% v/v on adry basis; (b) converting carbon monoxide to carbon dioxide bysubjecting the crude gas to a single stage of catalytic shift reactionto produce a raw gas having a carbon monoxide content of less than 0.5%v/v on a dry basis and in which at least 90% v/v of the total mediumboiling gas is nitrogen; (c) removing carbon dioxide and medium boilinggas, including the excess of nitrogen, from the raw gas by a pressureswing adsorption process, to give a product gas; and (d) methanating theproduct gas to convert residual of carbon oxides therein to methane; and(B) passing the ammonia synthesis gas over an ammonia synthesis catalystto product a reacted gas stream containing synthesized ammonia, andseparting synthesized ammonia from said reacted gas stream.
 2. A processaccording to claim 1, wherein at least 95% of the total medium boilinggas in the raw is nitrogen.
 3. A process according to claim 1, whereinthe raw gas has a total medium boiling gas to carbon dioxide volumeratio between 1.3 and 2.5.
 4. A process according to claim 1, whereinthe raw gas has a hydrogen to total medium boiling gas volume ratiobetween 1.5 an 2.1.
 5. A process according to claim 1, wherein thecatalytic shift reaction is operated in indirect heat exchange with acoolant, thereby producing a stream of heated coolant.
 6. A processaccording to claim 5, wherein the catalytic shift reaction is controlledat 230° to 280° C.
 7. A process according to claim 6, wherein the shiftoutlet temperature is 10° to 30° C. lower than the inlet temperature. 8.A process according to claim 5, wherein the coolant is water underpressure.
 9. A process according to claim 8, wherein the coolant isboiling water.
 10. A process according to claim 8, wherein:(a) the wastegas from the pressure swing adsorption, containing carbon dioxide andcombustibles, is used to power a gas expander producing at least part ofthe power required to compress at least one of the ammonia synthesis gasproduct, said carbonaceous feedstock, steam, and gas containing oxygenand nitrogen; and (b) the heat exchange in the catalytic shift reactionis operated without producing external steam.
 11. A process according toclaim 5, wherein in step (a) using a volatile hydrocarbon as thefeedstock, is reacted with steam over a externally heated catalyst in aprimary reformer and then the resultant primary reformed gas stream isreacted with the gas containing oxygen and nitrogen and passed over acatalyst in a secondary reformer so as to produce a hot secondaryreformed gas stream, and the heat required for the primary reforming isobtained by indirect heat exchange with the hot secondary reformed gasstream.
 12. A process according to claim 1, in which the pressure swingadsorption system used includes an adsorption step producing a productgas varying in composition with time between an initial and finallyrelatively low hydrogen/nitrogen volume ratio and an intermediaterelatively high hydrogen/nitrogen ratio, the integratedhydrogen/nitrogen ratio over the whole adsorption step being at a levelrequired in the product ammonia synthesis gas.
 13. A process accordingto claim 12, in which the adsorption pressure is in the range 30 to 40bar abs and is 10 to 15 times the purge and final dump pressure.
 14. Aprocess according to claim 12, wherein the pressure swing adsorptionstep the raw gas is fed simultaneously to a plurality of adsorbent bedsout of phase with one another.
 15. A process according to claim 12,wherein the pressure swing adsorption step includes a step wherein eachbed is repressurized counter-currently with the product gas prior to itsadsorption duty.
 16. A process for the production of ammoniacomprising(A) forming ammonia synthesis gas by(a) preheating a reactantsstream containing a volatile hydrocarbon feedstock and steam; (b)reacting said hydrocarbon feedstock with said steam over a catalyst in aprimary reformer to produce a primary reformed gas stream containingunreacted steam; (c) reacting said primary reformed gas stream with agas containing oxygen and nitrogen and passing the resultant mixtureover a catalyst in a secondary reformer so as to produce a hot secondaryreformed gas stream; the reforming process conditions and reactantproportions being such that the secondary reformed gas streamcontains:unreacted steam; hydrogen; carbon dioxide; and medium boilinggas consisting of:carbon monoxide; nitrogen in an excess of thatrequired in the ammonia synthesis gas; methane; and, optionally, argon;(d) cooling the secondary reformed gas stream without the production ofexternal steam by indirect heat exchange with(i) the reactants streamand (ii) a stream of heated water comprising:hot water, or a mixture ofhot water and steam; thereby forming a cooled secondary reformed gasstream; effecting the preheating of the reactants stream; and forming astream of further heated water; (e) converting carbon monoxide to carbondioxide by subjecting the cooled secondary reformed gas stream to asingle stage of catalytic shift reaction to produce a raw gas having acarbon monoxide content of less than 0.5% v/v on a dry basis, the shiftprocess being operated in heat exchange with a stream of water underpressure, thereby producing the stream of heated water employed in step(d); (f) contacting the volatile hydrocarbon feedstock, in the gaseousstate, with the stream of further heated water produced in step (d)prior to the preheating of the reactants stream in step (d), whereby atleast part of the steam required in the reactants stream is introduced;(g) removing carbon dioxide and medium boiling gas, including the excessof nitrogen, from the raw gas by a pressure swing adsorption process, togive a produce gas; and (h) methanating the product gas to convertresidual carbon oxides therein to methane; and (B) passing the ammoniasynthesis gas over an ammonia synthesis catalyst to produce a reactedgas stream containing synthesized ammonia, and separating synthesizedammonia from said reacted gas stream.
 17. A process according to claim16 wherein, prior to the indirect heat exchange of the secondaryreformed gas stream with the reactants stream and the stream of hotwater, the secondary reformed gas stream is partially cooled by indirectheat with the reactants stream while the latter is undergoing theprimary reforming step (b) thereby supplying the heat required for theprimary reforming step.
 18. A process for the production of ammoniacomprising(A) forming ammonia synthesis gas by(a) preheating a reactantsstream containing a volatile hydrocarbon feedstock and steam; (b)reacting said hydrocarbon feedstock with said steam over a catalyst in aprimary reformer to produce a primary reformed gas stream containingunreacted steam; (c) reacting said primary reformed gas stream with agas containing oxygen and nitrogen and passing the resultant mixtureover a catalyst in a secondary reformer so as to produce a hot secondaryreformed gas stream; the reforming process conditions and reactantproportions being such that the secondary reformed gas streamcontains:unreacted steam; hydrogen; carbon dioxide; and medium boilinggas consisting of:carbon monoxide; nitrogen in an excess of thatrequired in the ammonia synthesis gas; methane; and, optionally, argon;(d) cooling the secondary reformed gas stream without the production ofexternal steam by indirect heat exchange with(i) the reactants streamand (ii) a stream of heated water comprising:hot water, or a mixture ofhot water and steam; thereby forming a cooled secondary reformed gasstream; effecting the preheating of the reactants stream; and forming astream of further heated water; (e) converting carbon monoxide to carbondioxide by subjecting the cooled secondary reformed gas stream to asingle stage of catalytic shift reaction to produce a raw gas having acarbon monoxide content of less than 0.5% v/v on a dry basis, the shiftprocess being operated in heat exchange with a stream of coolant,thereby producing a stream of heated coolant; (f) cooling the stream ofheated coolant by indirect heat exchange with a stream of water, therebyproducing the stream of heated water employed in step (d); (g)contacting the volatile hydrocarbon feedstock, in the gaseous state,with the stream of further heated water produced in step (d) prior tothe preheating of the reactants stream in step (d), whereby at leastpart of the steam required in the reactants stream is introduced; (h)removing carbon dioxide and medium boiling gas, including the excess ofnitrogen, from the raw gas by a pressure swing adsorption process, togive a produce gas; and (i) methanating the product gas to convertresidual carbon oxides therein to methane; and (B) passing the ammoniasynthesis gas over an ammonia synthesis catalyst to produce a reactedgas stream containing synthesized ammonia, and separating synthesizedammonia from said reacted gas stream.